This invention relates to a process for the separation of a gas containing hydrocarbons.
Ethylene, ethane, propylene, propane, and/or heavier hydrocarbons can be recovered from a variety of gases, such as natural gas, refinery gas, and synthetic gas streams obtained from other hydrocarbon materials such as coal, crude oil, naphtha, oil shale, tar sands, and lignite. Natural gas usually has a major proportion of methane and ethane, i.e., methane and ethane together comprise at least 50 mole percent of the gas. The gas also contains relatively lesser amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as hydrogen, nitrogen, carbon dioxide, and other gases.
The present invention is generally concerned with the recovery of ethylene, ethane, propylene, propane, and heavier hydrocarbons from such gas streams. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 91.6% methane, 4.2% ethane and other C2 components, 1.3% propane and other C3 components, 0.4% iso-butane, 0.3% normal butane, 0.5% pentanes plus, 1.4% carbon dioxide, with the balance made up of nitrogen. Sulfur containing gases are also sometimes present.
The historically cyclic fluctuations in the prices of both natural gas and its natural gas liquid (NGL) constituents have at times reduced the incremental value of ethane, ethylene, propane, propylene, and heavier components as liquid products. This has resulted in a demand for processes that can provide more efficient recoveries of these products, for processes that can provide efficient recoveries with lower capital investment and lower operating costs, and for processes that can be easily adapted or adjusted to vary the recovery of a specific component over a broad range. Available processes for separating these materials include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.
The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,712,880; 6,915,662; reissue U.S. Pat. No. 33,408; U.S. Application Publ. No. 2002/0166336 A1; and co-pending application Ser. No. 11/201,358 describe relevant processes (although the description of the present invention in some cases is based on different processing conditions than those described in the cited patents and applications).
In a typical cryogenic expansion recovery process, a feed gas stream under pressure is cooled by heat exchange with other streams of the process and/or external sources of refrigeration such as a propane compression-refrigeration system. As the gas is cooled, liquids may be condensed and collected in one or more separators as high-pressure liquids containing some of the desired C2+ or C3+ components. Depending on the richness of the gas and the amount of liquids formed, the high-pressure liquids may be expanded to a lower pressure and fractionated. The vaporization occurring during expansion of the liquids results in further cooling of the stream. Under some conditions, pre-cooling the high pressure liquids prior to the expansion may be desirable in order to further lower the temperature resulting from the expansion. The expanded stream, comprising a mixture of liquid and vapor, is fractionated in a distillation (demethanizer or deethanizer) column. In the column, the expansion cooled stream(s) is (are) distilled to separate residual methane, nitrogen, and other volatile gases as overhead vapor from the desired C2 components, C3 components, and heavier hydrocarbon components as bottom liquid product, or to separate residual methane, C2 components, nitrogen, and other volatile gases as overhead vapor from the desired C3 components and heavier hydrocarbon components as bottom liquid product.
If the feed gas is not totally condensed (typically it is not), the vapor remaining from the partial condensation can be passed through a work expansion machine or engine, or an expansion valve, to a lower pressure at which additional liquids are condensed as a result of further cooling of the stream. The pressure after expansion is essentially the same as the pressure at which the distillation column is operated. The expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams. The vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed.
In the ideal operation of such a separation process, the residue gas leaving the process will contain substantially all of the methane in the feed gas with essentially none of the heavier hydrocarbon components and the bottoms fraction leaving the demethanizer will contain substantially all of the heavier hydrocarbon components with essentially no methane or more volatile components. In practice, however, this ideal situation is not obtained for two main reasons. The first reason is that the conventional demethanizer is operated largely as a stripping column. The methane product of the process, therefore, typically comprises vapors leaving the top fractionation stage of the column, together with vapors not subjected to any rectification step. Considerable losses of C2, C3, and C4+ components occur because the top liquid feed contains substantial quantities of these components, resulting in corresponding equilibrium quantities of C2 components, C3 components, C4 components, and heavier hydrocarbon components in the vapors leaving the top fractionation stage of the demethanizer. The loss of these desirable components could be significantly reduced if the rising vapors could be brought into contact with a significant quantity of liquid (reflux) capable of absorbing the C2 components, C3 components, C4 components, and heavier hydrocarbon components from the vapors.
The second reason that this ideal situation cannot be obtained is that carbon dioxide contained in the feed gas fractionates in the demethanizer and can build up to concentrations of as much as 5% to 10% or more in the tower even when the feed gas contains less than 1% carbon dioxide. At such high concentrations, formation of solid carbon dioxide can occur depending on temperatures, pressures, and the liquid solubility. It is well known that natural gas streams usually contain carbon dioxide, sometimes in substantial amounts. If the carbon dioxide concentration in the feed gas is high enough, it becomes impossible to process the feed gas as desired due to blockage of the process equipment with solid carbon dioxide (unless carbon dioxide removal equipment is added, which would increase capital cost substantially). The present invention provides a means for generating a supplemental liquid reflux stream that will improve the recovery efficiency for the desired products while simultaneously substantially mitigating the problem of carbon dioxide icing.
In recent years, the preferred processes for hydrocarbon separation use an upper absorber section to provide additional rectification of the rising vapors. The source of the reflux stream for the upper rectification section is typically a recycled stream of residue gas supplied under pressure. The recycled residue gas stream is usually cooled to substantial condensation by heat exchange with other process streams, e.g., the cold fractionation tower overhead. The resulting substantially condensed stream is then expanded through an appropriate expansion device, such as an expansion valve, to the pressure at which the demethanizer is operated. During expansion, a portion of the liquid will usually vaporize, resulting in cooling of the total stream. The flash expanded stream is then supplied as top feed to the demethanizer. Typically, the vapor portion of the expanded stream and the demethanizer overhead vapor combine in an upper separator section in the fractionation tower as residual methane product gas. Alternatively, the cooled and expanded stream may be supplied to a separator to provide vapor and liquid streams, so that thereafter the vapor is combined with the tower overhead and the liquid is supplied to the column as a top column feed. Typical process schemes of this type are disclosed in U.S. Pat. Nos. 4,889,545; 5,568,737; 5,881,569; 6,712,880; and in Mowrey, E. Ross, “Efficient, High Recovery of Liquids from Natural Gas Utilizing a High Pressure Absorber”, Proceedings of the Eighty-First Annual Convention of the Gas Processors Association, Dallas, Tex., Mar. 11-13, 2002.
The present invention also employs an upper rectification section (or a separate rectification column in some embodiments). However, two reflux streams are provided for this rectification section. The upper reflux stream is a recycled stream of residue gas as described above. In addition, however, a supplemental reflux stream is provided at a lower feed point by using a side draw of the vapors rising in a lower portion of the tower (which may be combined with some of the separator liquids). Because of the relatively high concentration of C2 components and heavier components in the vapors lower in the tower, a significant quantity of liquid can be condensed in this side draw stream without elevating its pressure, often using only the refrigeration available in the cold vapor leaving the upper rectification section. This condensed liquid, which is predominantly liquid methane and ethane, can then be used to absorb C3 components, C4 components, and heavier hydrocarbon components from the vapors rising through the lower portion of the upper rectification section and thereby capture these valuable components in the bottom liquid product from the demethanizer. Since the lower reflux stream captures essentially all of the C3+ components, only a relatively small flow rate of liquid in the upper reflux stream is needed to absorb the C2 components remaining in the rising vapors and likewise capture these C2 components in the bottom liquid product from the demethanizer.
Heretofore, such a vapor side draw feature has been employed in C3+ recovery systems, as illustrated in the assignee's U.S. Pat. No. 5,799,507. The process and apparatus of U.S. Pat. No. 5,799,507, however, are unsuitable for high ethane recovery. Surprisingly, applicants have found that C2 recoveries may be improved without sacrificing C3+ component recovery levels or system efficiency by combining the side draw feature of the assignee's U.S. Pat. No. 5,799,507 invention with the residue reflux feature of the assignee's U.S. Pat. No. 5,568,737.
In accordance with the present invention, it has been found that C2 component recoveries in excess of 97 percent can be obtained with no loss in C3+ component recovery. The present invention provides the further advantage of being easily adapted to using much of the equipment required to implement assignee's U.S. Pat. No. 5,799,507, resulting in lower capital investment costs compared to other prior art processes. In addition, the present invention makes possible essentially 100 percent separation of methane and lighter components from the C2 components and heavier components while maintaining the same recovery levels as the prior art and improving the safety factor with respect to the danger of carbon dioxide icing. The present invention, although applicable at lower pressures and warmer temperatures, is particularly advantageous when processing feed gases in the range of 400 to 1500 psia [2,758 to 10,342 kPa(a)] or higher under conditions requiring NGL recovery column overhead temperatures of −50° F. [−46° C.] or colder.
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d′Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
FIG. 1 is a process flow diagram showing the design of a processing plant to recover C3+ components from natural gas using prior art according to assignee's U.S. Pat. No. 5,799,507. In this simulation of the process, inlet gas enters the plant at 120° F. [49° C.] and 1040 psia [7,171 kPa(a)] as stream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.
The feed stream 31 is cooled in heat exchanger 10 by heat exchange with cool residue gas at −88° F. [−67° C.] (stream 52) and flash expanded separator liquids (stream 33a). The cooled stream 31a enters separator 11 at −34° F. [−37° C.] and 1025 psia [7,067 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33). The separator liquid (stream 33) is expanded to slightly above the operating pressure of fractionation tower 19 by expansion valve 12, cooling stream 33a to −67° F. [−55° C.]. Stream 33a enters heat exchanger 10 to supply cooling to the feed gas as described previously, heating stream 33b to 116° F. [47° C.] before it is supplied to fractionation tower 19 at a lower mid-column feed point.
The separator vapor (stream 32) enters a work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 17 expands the vapor substantially isentropically to the tower operating pressure of approximately 420 psia [2,894 kPa(a)], with the work expansion cooling the expanded stream 32a to a temperature of approximately −108° F. [−78° C.]. The typical commercially available expanders are capable of recovering on the order of 80-88% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 18) that can be used to re-compress the residue gas (stream 52a), for example. The partially condensed expanded stream 32a is thereafter supplied as feed to fractionation tower 19 at an upper mid-column feed point.
The deethanizer in tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. The deethanizer tower consists of two sections: an upper absorbing (rectification) section 19a that contains the trays and/or packing to provide the necessary contact between the vapor portion of the expanded stream 32a rising upward and cold liquid falling downward to condense and absorb the C3 components and heavier components; and a lower, stripping section 19b that contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The deethanizing section 19b also includes at least one reboiler (such as reboiler 20) which heats and vaporizes a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41, of methane, C2 components, and lighter components. Stream 32a enters deethanizer 19 at an upper mid-column feed position located in the lower region of absorbing section 19a of deethanizer 19. The liquid portion of expanded stream 32a commingles with liquids falling downward from the absorbing section 19a and the combined liquid continues downward into the stripping section 19b of deethanizer 19. The vapor portion of expanded stream 32a rises upward through absorbing section 19a and is contacted with cold liquid falling downward to condense and absorb the C3 components and heavier components.
A portion of the distillation vapor (stream 42) is withdrawn from the upper region of stripping section 19b. This stream is then cooled and partially condensed (stream 42a) in exchanger 22 by heat exchange with cold deethanizer overhead stream 38 which exits the top of deethanizer 19 at −114° F. [−81° C.] and with a portion of the cold distillation liquid (stream 47) withdrawn from the lower region of absorbing section 19a at −112° F. [−80° C.]. The cold deethanizer overhead stream is warmed to approximately −87° F. [−66° C.] (stream 38a) and the distillation liquid is heated to −43° F. [−42° C.] (stream 47a) as they cool stream 42 from −39° F. [−40° C.] to about −109° F. [−78° C.] (stream 42a). The heated and partially vaporized distillation liquid (stream 47a) is then returned to deethanizer 19 at a mid-point of stripping section 19b. 
The operating pressure in reflux separator 23 is maintained slightly below the operating pressure of deethanizer 19. This pressure difference provides the driving force that allows distillation vapor stream 42 to flow through heat exchanger 22 and thence into the reflux separator 23 wherein the condensed liquid (stream 44) is separated from the uncondensed vapor (stream 43). The uncondensed vapor stream 43 combines with the warmed deethanizer overhead stream 38a from exchanger 22 to form cool residue gas stream 52 at −88° F. [−67° C.].
The liquid stream 44 from reflux separator 23 is pumped by pump 24 to a pressure slightly above the operating pressure of deethanizer 19. The resulting stream 44a is then divided into two portions. The first portion (stream 45) is supplied as cold top column feed (reflux) to the upper region of absorbing section 19a of deethanizer 19. This cold liquid causes an absorption cooling effect to occur inside the absorbing (rectification) section 19a of deethanizer 19, wherein the saturation of the vapors rising upward through the tower by vaporization of liquid methane and ethane contained in stream 45 provides refrigeration to the section. Note that, as a result, both the vapor leaving the upper region (overhead stream 38) and the liquids leaving the lower region (liquid distillation stream 47) of absorbing section 19a are colder than the either of the feed streams (streams 45 and stream 32a) to absorbing section 19a. This absorption cooling effect allows the tower overhead (stream 38) to provide the cooling needed in heat exchanger 22 to partially condense the vapor distillation stream (stream 42) without operating stripping section 19b at a pressure significantly higher than that of absorbing section 19a. This absorption cooling effect also facilitates reflux stream 45 condensing and absorbing the C3 components and heavier components in the distillation vapor flowing upward through absorbing section 19a. The second portion (stream 46) of pumped stream 44a is supplied to the upper region of stripping section 19b of deethanizer 19 where the cold liquid acts as reflux to absorb and condense the C3 components and heavier components flowing upward from below so that vapor distillation stream 42 contains minimal quantities of these components.
In stripping section 19b of deethanizer 19, the feed streams are stripped of their methane and C2 components. The resulting liquid product stream 41 exits the bottom of deethanizer 19 at 225° F. [107° C.] (based on a typical specification of a ethane to propane ratio of 0.025:1 on a molar basis in the bottom product) before flowing to storage.
The cool residue gas (stream 52) passes countercurrently to the incoming feed gas in heat exchanger 10 where it is heated to 115° F. [46° C.] (stream 52a). The residue gas is then re-compressed in two stages. The first stage is compressor 18 driven by expansion machine 17. The second stage is compressor 25 driven by a supplemental power source which compresses the residue gas (stream 52c) to sales line pressure. After cooling to 120° F. [49° C.] in discharge cooler 26, the residue gas product (stream 52d) flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)], sufficient to meet line requirements (usually on the order of the inlet pressure).
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 1 is set forth in the following table:
TABLE I(FIG. 1)Stream Flow Summary—Lb. Moles/Hr [kg moles/Hr]StreamMethaneEthanePropaneButanes+C. DioxideTotal3125,3841,16136233240027,7143225,0851,10431518638927,15333299574714611561472,8371,0733271861694,595424,3471,7972612796,452431,2536900251,349443,0941,7282612545,103451,8871,0541611553,113461,207674100991,9903824,1311,0833037525,6655225,3841,1523040027,01441093593320700Recoveries*Propane99.08%Butanes+99.99%PowerResidue Gas Compression12,774 HP[21,000 kW]*(Based on un-rounded flow rates)
The FIG. 1 process is often the optimum choice for gas processing plants when recovery of C2 components is not desired, because it provides very efficient recovery of the C3+ components using equipment that requires less capital investment than other processes. However, the FIG. 1 process is not well suited to recovering C2 components, as C2 component recovery levels on the order of 40% are generally the highest that can be achieved without inordinate increases in the power requirements for the process. If higher C2 component recovery levels than this are desired, a different process is usually required, such as assignee's U.S. Pat. No. 5,568,737.
FIG. 2 is a process flow diagram showing one manner in which the design of the processing plant in FIG. 1 can be adapted to operate at a higher C2 component recovery level using a base case design according to assignee's U.S. Pat. No. 5,568,737. The process of FIG. 2 has been applied to the same feed gas composition and conditions as described previously for FIG. 1. However, in the simulation of the process of FIG. 2, certain equipment and piping have been added (shown by bold lines) while other equipment and piping have been removed from service (shown by light dashed lines) so that the process operating conditions can be adjusted to increase the recovery of C2 components to about 97%.
The feed stream 31 is cooled in heat exchanger 10 by heat exchange with a portion of the cool distillation column overhead stream (stream 48) at −15° F. [−26° C.], demethanizer liquids (stream 39) at −33° F. [−36° C.], demethanizer liquids (stream 40) at 37° F. [3° C.], and the pumped demethanizer bottoms liquid (stream 41a) at 60° F. [16° C.]. The cooled stream 31a enters separator 11 at 4° F. [−16° C.] and 1025 psia [7,067 kPa(a)] where the vapor (stream 32) is separated from the condensed liquid (stream 33).
The separator vapor (stream 32) is divided into two streams, 34 and 36. Stream 34, containing about 30% of the total vapor, is combined with the separator liquid (stream 33). The combined stream 35 passes through heat exchanger 22 in heat exchange relation with the cold distillation column overhead stream 38 where it is cooled to substantial condensation. The resulting substantially condensed stream 35a at −138° F. [−95° C.] is then flash expanded through expansion valve 16 to the operating pressure of fractionation tower 19, 412 psia [2,839 kPa(a)]. During expansion a portion of the stream is vaporized, resulting in cooling of the total stream. In the process illustrated in FIG. 2, the expanded stream 35b leaving expansion valve 16 reaches a temperature of −141° F. [−96° C.] and is supplied to fractionation tower 19 at an upper mid-column feed point.
The remaining 70% of the vapor from separator 11 (stream 36) enters a work expansion machine 17 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 17 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 36a to a temperature of approximately −80° F. [−62° C.]. The partially condensed expanded stream 36a is thereafter supplied as feed to fractionation tower 19 at a lower mid-column feed point.
The recompressed and cooled distillation stream 38e is divided into two streams. One portion, stream 52, is the residue gas product. The other portion, recycle stream 51, flows to heat exchanger 27 where it is cooled to −1° F. [−18° C.] (stream 51a) by heat exchange with a portion (stream 49) of cool distillation column overhead stream 38a at −15° F. [−26° C.]. The cooled recycle stream then flows to exchanger 22 where it is cooled to −138° F. [−95° C.] and substantially condensed by heat exchange with cold distillation stream 38. The substantially condensed stream 51b is then expanded through an appropriate expansion device, such as expansion valve 15, to the demethanizer operating pressure, resulting in cooling of the total stream. In the process illustrated in FIG. 2, the expanded stream 51c leaving expansion valve 15 reaches a temperature of −145° F. [−98° C.] and is supplied to the fractionation tower as the top column feed. The vapor portion (if any) of stream 51c combines with the vapors rising from the top fractionation stage of the column to form distillation stream 38, which is withdrawn from an upper region of the tower.
The demethanizer in tower 19 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing. As is often the case in natural gas processing plants, the fractionation tower may consist of two sections. The upper section 19a is a separator wherein the top feed is divided into its respective vapor and liquid portions, and wherein the vapor rising from the lower distillation or demethanizing section 19b is combined with the vapor portion (if any) of the top feed to form the cold demethanizer overhead vapor (stream 38) which exits the top of the tower at −142° F. [−97° C.]. The lower, demethanizing section 19b contains the trays and/or packing and provides the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section 19b also includes reboilers (such as trim reboiler 20 and the reboiler and side reboiler described previously) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41, of methane and lighter components.
The liquid product stream 41 exits the bottom of the tower at 55° F. [13° C.], based on a typical specification of a methane to ethane ratio of 0.025:1 on a molar basis in the bottom product. Pump 21 delivers stream 41a to heat exchanger 10 as described previously where it is heated to 116° F. [47° C.] before flowing to storage. The demethanizer overhead vapor stream 38 passes countercurrently to the incoming feed gas and recycle stream in heat exchanger 22 where it is heated to −15° F. [−26° C.]. The heated stream 38a is divided into two portions (streams 49 and 48), which are heated to 116° F. [47° C.] and 78° F. [25° C.], respectively, in heat exchanger 27 and heat exchanger 10. The heated streams recombine to form stream 38b at 84° F. [29° C.] which is then re-compressed in two stages, compressor 18 driven by expansion machine 17 and compressor 25 driven by a supplemental power source. After stream 38d is cooled to 120° F. [49° C.] in discharge cooler 26 to form stream 38e, recycle stream 51 is withdrawn as described earlier to form residue gas stream 52 which flows to the sales gas pipeline at 1040 psia [7,171 kPa(a)].
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 2 is set forth in the following table:
TABLE II(FIG. 2)Stream Flow Summary—Lb. Moles/Hr [kg moles/Hr]StreamMethaneEthanePropaneButanes+C. DioxideTotal3125,3841,16136233240027,7143225,3071,14534825239727,52433771614803190347,719349106771218,3953617,58879624217527619,129357,7963651201571248,5853829,587400014629,859514,231600214,2705225,356340012525,58941281,1273623322752,125Recoveries*Ethane97.04%Propane100.00%Butanes+100.00%PowerResidue Gas Compression14,219 HP[23,376 kW]*(Based on un-rounded flow rates)
By modifying the FIG. 1 equipment and piping as shown in FIG. 2, the natural gas processing plant can now achieve 97% recovery of the C2 components in the feed gas. This means that the plant has the flexibility to operate as shown in FIG. 2 and recover essentially all of the C2 components when the value of liquid C2 components is attractive, or to operate as shown in FIG. 1 and reject the C2 components to the plant residue gas when the C2 components are more valuable as gaseous fuel. However, the required modifications require much additional equipment and piping (as shown by the bold lines) and do not make use of much of the equipment present in the FIG. 1 plant (shown by the light dashed lines), so the capital cost of a plant designed to operate using both the FIG. 1 process and the FIG. 2 process will be higher than is desirable. (Note that although the FIG. 2 process can be adapted to reject the C2 components like the FIG. 1 process, the power consumption when operating in this manner is essentially the same as that shown in Table II. Since this is about 11% higher than that of the FIG. 1 process as shown in Table I, the operating cost of a plant using the FIG. 1 process is considerably lower than that of one using the FIG. 2 process in this manner.)